Multiple-stage cascade conversion of black oil



Nov. 5, 196s -w K. r GLEI'M ET AL 3,409,538v

MULTIPLE-STAGE CASCADE CONVERSION OF' BLACK OIL Filed Appil 24, 1967 T TOR/VEYS Wwf' United States Patent O 3,409,53 MULTIPLE-STAGE CASCADE CONVERSION F BLACK OIL William K. T. Gleim, Island Lake, and Mark J. OHara,

Prospect Heights, Ill., assignors to Universal Oil Products Company, Des Plaines, Ill., a corporation of Delaware Filed Apr. 24, 1967, Ser. No. 633,052 1 Claim. (Cl. 208-59) ABSTRACT OF THE DISCLOSURE A process for the conversion of heavy, contaminated hydrocarbonaceous charge stocks commonly referred to in the art as black oils. The process constitutes successive stages of conversion, each of which involves a catalytic reaction zone .coupled with a hot ash separation zone. A preferred processing technique encompasses successively higher average catalyst temperatures in the succeeding conversion zones.

Applicability of invention The invention described herein is applicable to a process for the conversion of contaminated petroleum crude oil, as well as the heavier fractions which may be derived therefrom, into lower boiling hydrocarbon products. More specifically, the present invention is directed toward a process for converting atmospheric tower bottoms products, vacuum tower bottoms products (vacuum residuum), crude oil residuum, topped crude oils, crude oils extracted from tar sands, etc., all of which are commonly referred to as black oils.

Black oils contain high molecular weight sulfurous compounds, nitrogenous compounds, both in relatively large quantities, high molecular weight organo-metallic complexes comprising nickel and vanadium as the metallic component, and asphaltenic material which is insoluble in normally liquid hydrocarbons such as pentane and heptane. 'The asphaltics are generally found to be complexed with, or linked to sulfur and, to a certain extent, with metallic contaminants. A black oil is often characterized as a heavy hydrocarbonaceous material of which more than 10.0% (by volume) boils above a temperature of 1050 F., and as having a gravity, API at 60 F., of less than 1.0% by weight, and are often in excess of 3.0% by weight. Conradson Carbon Residue factors exceed 1.0 weight percent, and a great proportion of black oils exhibit a Conradson Carbon Residue factor above 10.0. The utilization of these highly contaminated black oils as a source of more valuable liquid hydrocarbon products is precluded unless the sulfur, nitrogen and asphaltene content is sharply reduced, and a significant proportion of the material can be converted into distillable hydrocarbons-ie. those boiling below about 1050 F. (as determined by ASTM Method D-ll60).

The process encompassed by the present invention is particularly directed toward the catalytic conversion of black oils into distillable hydrocarbons. Specic examples of the crude oils to which the present scheme is uniquely adaptable, include a vacuum tower bottoms product having a gravity of 7.1 API at 60 F., and containing 4.05% by weight of sulfur and 23.7% by weight of asphalts; a topped Middle East Kuwait crude oil, having a gravity of 11.0 API and containing 10.1% by weight of asphaltenes and 5.2% by weight of sulfur; a vacuum residuum having a gravity of 8.8 API and containing 3.0% sulfur and 4300 p.p.m. (by weight) of nitrogen; vacuum bottoms having a gravity of 5.4 API, and containing 6.15% sulfur, 233 p.p.m. (by weight) of metals and 12.8% by 3,409,538 Patented Nov. 5, 1968 weight of heptane-insoluble asphaltenic material; and, a reduced crude having a gravity of 11.5 API, containing 4.2% sulfur, 3400 p.p.m. nitrogen, 166 p.p.m. of metals and 8.6% by weight of heptane-insolubles.

The present invention affords the conversion of up to 80.0% by volume of such material into distillable hydrocarbons, heretofore having been considered impossible to achieve, especially on an economically feasible basis. The principal diiicnlty resides in the lack of sulfur stability of many catalytic composites employed in current processes, and primarily from the presence of large quantities of asphaltic material and other non-distillables. Generally, the asphaltenic material is found to be dispersed within the black oil, and when subjected to heat, as in a vacuum distillation process, has the tendency to ilocculate and polymerize whereby the conversion thereof to more valuable oil-soluble products becomes extremely difficult. Thus, the heavy bottoms from a crude oil vacuum distillation column (vacuum residuum), indicates a Conradson Carbon Residue factor of 16.0% by weight.

Prior art It must be acknowledged that published literature recognizes various types of processes designed to effect the hydrorening of black oils. Thus, many literature references and/ or publications might be found which disclose propane deasphalting followed by thermal cracking of coking of the resulting normally liquid product, desalting followed by halogen hydride treatment, to coagulate the metallic-containing asphaltenes, etc. It is noteworthy that the latter processing schemes are unconcerned with catalytic processing of black oils.

Furthermore, with respect to the area of catalytic processing, two principal approaches have been advanced: liquid-phase hydrogenation and vapor-phase hydrocracking. In the former, liquid phase oil is passed (generally upwardly), in admixture with hydrogen, into a fixedfluidized bed of catalyst particles. Although perhaps effective in converting at least a portion of the oil-soluble metallic complexes, this type process is relatively ineifective with respect to asphaltenes dispersed in the charge, since the probability of effecting simultaneous contact between the catalyst particle and the asphaltenic molecule is remote. Furthermore, since the hydrogenation reaction zone is generally maintained at an elevated temperature of at least about 500 C. (932 R), the retention of unconverted asphaltenes suspended in a free liquid phase oil for an extended period of time, results in further agglomeration, making conversion thereof substantially more difcult. The eiciency of hydrogen to oil contact obtainable by bubbling hydrogen through an extensive liquid body is relatively low. Some processes have been described which rely extensively upon thermal cracking in the presence of hydrogen; such processes are unable to effect the conversion of asphaltenic material. Briefly, the present invention is directed toward a process whereby the asphaltenic material is maintained in a dispersed state in a liquid phase rich in hydrogen. This material comes into intimate contact with a fixed-bed of catalyst capable of effecting reaction between the hydrogen and the asphaltenes; the liquid phase is itself dispersed in a hydrogen-rich gas phase so that the dissolved hydrogen is continually replenished. The two-fold dispersion and rapid, intimate contacting with the catalytic surface overcomes the difculties encountered in previous processes whereby excessive residence times and depletion of localized hydrogen supply permit agglomeration of asphaltenes and other high molecular weight species. Such agglomerates are even less availablevto hydrogen and are not, therefore, susceptible to catalytic reaction. They eventually form coke which becomes deposited on the catalyst, thereby further reducing catalytic activity within the system.

An object of the present invention is to provide a conversion process to effect the hydrorefining/hydrocracking of contaminated hydrocarbonaceous material. A corollary objective is to afford a novel multiple-stage process for the conversion of asphaltene-containing black oils.

Specifically, it is an object of our invention to provide an efficient, economical process for maximizing the production of gasoline and middle-distillate boiling range hydrocarbons from a black oil boiling substantially completely above the middle-distillate boiling range.

These objects, and others, are achieved through the practice of the present invention as hereinafter set forth in greater detail.

Summary of invention In one embodiment, the present invention encompasses a process for the conversion of a hydrocarbonaceous black oil, containing hydrocarbons boiling above about 1050 F., into lower-boiling distillable products, which process comprises the steps of: (a) admixing said black oil with hydrogen and passing said mixture, at a temperature above about 500 F., into a first catalytic conversion zone maintained at a pressure above about 1000 p.s.i.g.; (b) separating the resulting first conversion zone efiluent in a first separation zone at a pressure of from subatmospheric to about 200 p.s.i.g., to provide a first vapor phase and a first liquid phase; (c) introducing said first liquid phase into a second catalytic conversion zone, at a temperature above about 600 F., and a pressure above about 1000 p.s.i.g.; (d) separating the resulting second conversion zone eflluent in a second separation zone at a pressure of from subatmospheric to about 100 p.s.i.g., to provide a second vapor phase and a second liquid phase; (e) introducing said second liquid phase into a third catalytic conversion zone, at a temperature above about 775 F., and a pressure above about 1000 p.s.i.g.; (f) separating the resulting third conversion zone effluent in a third separation zone at a pressure of from subatmospheric to about 100 p.s.i.g., to provide a third vapor phase and a third liquid phase containing hydrocarbonaceous material boiling above about 1050 F.; (g) combining said first, second and third vapor phases with hydrogen, and introducing the resulting mixture into a fourth catalytic conversion zone at a temperature of from about 500 F. to about 850 F. and a pressure of from about 1000 to about 4000 p.s.i.g.; and, (h) separating the resulting fourth conversion zone eluent in a fourth separation zone at a temperature of from about 60 F. to about 140 F. and a pressure substantially the same as imposed upon said fourth conversion zone, to provide a fourth liquid phase comprising distillable hydrocarbons and a fourth vapor phase rich in hydrogen.

The process as described in the foregoing embodiment may be further characterized in that the inlet temperature of the first catalytic conversion zone is less than the inlet temperature of the second conversion zone, in turn less than that of the third conversion zone, Preferably, the first conversion zone will function at an inlet temperature of from about 500 F. to about 600 F.; the second conversion zone at an inlet temperature of from 600 F. t0 about 775 F.; and, the third zone at an inlet temperature of about 775 F. to about 850 F. During operation of the process, the temperature control point is the inlet to the catalyst bed disposed within the zone, and the temperature to which the hydrocarbonaceous charge to a particular zone is raised is adjusted accordingly. Thus, it is understood that 'reference to a conversion zone temperature is intended to connote the temperature at the initial portion of the catalyst bed.

An essential feature of our invention resides in the use of these particular temperature ranges, the precise temperature in any Situation being primarily dependent upon the physical and chemical characteristics of the charge stock, as well as the desired product distribution. As hereinafter indicated, the present process possesses an inherent degree of flexibility with respect to the comparative quantities of gasoline and middle-distillate hydrocarbons produced. The novel procedure involving the conversion zones and the accompanying flash zones assures the fresh (or unconverted) asphaltenes of dilution in a high boiling aromatic system afforded by the liquid phases from the flash zones. Having been subjected to conversion conditions at least once, the aromatics are sufficiently more thermal stable than the asphaltenes, with the result that such dilution facilitates conversion of the latter. In a typical, currently-practiced technique, the asphaltenes are simultaneously subjected to the higher temperatures, and,I being in competition with other constituents for active catalyst sites, would preferentially react thermally to cause coke laydown. As hereinafter indicated in a specific example, our inventive concept arose from the recognition that the charge stock is to be preheated in the presence of a catalyst at a temperature insufficiently high to contribute to significant degree of hydrocracking. This operation is compared to a typical, standard operation in which the charge stock was heated to reaction temperature in contact with an inert material-Le. quartz chips.

Other embodiments of our invention reside in particular operating conditions and the use of specific internal recycle streams. The latter include recycle of the hydrogen-rich fourth gaseous phase from the cold separator to combine with the first and second liquid phases and the third vaporous phase prior to reacting the same. At least a part of the first liquid phase may be diverted and combined with the heated charge and make-up hydrogen mixture, and serves as a solvent stream to keep the asphaltenes dispersed and available to both hydrogen and catalyst in the reaction zone. In another embodiment, a portion of the first liquid phase may be cooled and recycled to the inlet of the first separation zone to serve as a quench of the reaction zone effluent such that the temperature in said first separation zone is maintained below a maximum level of 650 F. When processing variables demand, that portion of the first liquid phase being diverted from the stream introduced into the second reaction zone may be combined en toto with the fresh charge stock and make-up hydrogen, en toto with the first reaction zone effluent, or in part with each of these two streams. As hereinafter discussed with reference to the embodiment illustrated in the accompanying drawing, the first, second and third vapor phases are combined as the charge to a fourth reaction zone. This zone serves principally as a hydro-cracking zone for the greater degree of conversion of high boiling hydrocarbons into lower boiling products, primarily gasoline, kerosene and midle-distillates. Provision may be made to recycle a portion of the heavier fraction-ie. from 525 F. to about 650 F.-to combine with the fresh charge stock and provide dilution for the asphaltenes therein.

A hot flash separation zone functions at a significantly reduced pressure of from subatmospheric to about 200 p.s.i.g., and may consist of a low-pressure flash zonei.e. about 60 p.s.i.g.-in combination with a vacuum column maintained at about 50-60 mm. of Hg absolute. The succeeding hot flash system following each conversion zone serves to eliminate the difficulties stemming from emulsification problems by ultimately providing a residuum fraction containing the unconverted asphaltenes and a significant amount of those sulfurous compounds not converted in the preceding reaction zones. Furthermore, subsequent separations and/ or distillations are greatly simplified.

Before describing my invention with reference to the accompanying drawing, several definitions are believed necessary in order that a clear understanding be afforded.

In the present specification and the appended claims, the phrase pressure substantially the same as, is intended to connote that pressure under which a succeeding vessel is maintained, allowing only for the pressure drop experienced as a result of the flow of fluids through the system. That is, no specic, intentional means will be employed to reduce the pressure. Similarly, unless otherwise specified, the phrase temperature substantially the same as, is used to indicate that the only reduction in temperature stems from normally experienced lossdue to the ow of material from one piece of equipment to another, or from conversion of sensible to latent heat by ashing where a pressure drop occurs.

The phrase, hydrocarbons boiling within the gasoline boiling range, or gasoline boiling range hydrocarbons, is intended to connote those hydrocarbons boiling at temperatures up to about 400 F. including C5-hydrocarbons, and, as in some localities, C4-hydrocarbons. Although the end boiling point of gasoline is sometimes considered to be as high as 425 F., or even 450 F., the use of the term herein will allude to a hydrocarbon fraction having a nominal end boiling point of about 400 F. Middle-distillate hydrocarbons, whether inclusive of kerosene, will connote a hydrocarbon mixture boiling above 400 F. and having a nominal end boiling point in the range of 650 F. to about 700 F.

Likewise, a black oil is intended to connote a hydrocarbonaceous mixture of which at least about 10.0% boils above a temperature of about l050 F., and which has a gravity, API at 60 F., of 20.0 or less. The greater proportion of such black oils contains 60.0% or more of material boiling above 1050 F., and in many instances, the material is considered totally non-distillable. Distillable hydrocarbons are those normally liquid hydrocarbons, including pentanes, having boiling points below about 1050 F. Conversion conditions are intended to be those conditions imposed upon the conversion zones in order to convert a substantial portion of the black oil to distillable hydrocarbons. As will be readily noted by those skilled in the art of petroleum rening techniques, the conversion conditions hereinafter enumerated are significantly less severe than those being curently commercially employed. The distinct economic advantages, over and above those inherent in producing the more valuable distillable hydrocarbons, will be imrnediately recognized. Since the bulk of the reactions being effected are exothermic, the temperature increases through the catalyst bed, and the reaction zone etiluent will be at a temperature higher than at the inlet to the catalyst bed. In order that catalyst stability be preserved, it is preferred to control the inlet Itemperature such that the efuent temperature from any one of the reaction zones does not exceed about 950 F. Hydrogen is mixed with the black oil charge stock, by means of compressive recycle, in an amount of from about 1,000 to about 50,000 s.c.f./|bbl. at the selected operating pressure, and preferably in an amount of from about 1,000 to about 10,000 s.c.f./ bbl. The operating pressure will be greater than 1,000 p.s.i.g., and generally in the range of about 1,500 p.s.i.g. to about 4,000 p.s.i.g. The black oil passes through the catalyst at a liquid hourly space velocity (defined as volumes of liquid hydrocarbon charge per hour, measured at 60 F. per volume of catalyst disposed in the reaction zone) of from about 0.25 to about 5.0. It is particularly preferred to introduce the mixture of black oil and hydrogen into the vessel in such a manner that the same passes through the vessel in downward flow. The internals of the vessel may be constructed in any suitable manner capable of providing the required intimate contact between the liquid charge stock, the gaseous mixture and the catalyst. In some instances, it may be desirable to provide the reaction zone with a packed bed of inert material such as particles of granite, porcelain, berl saddles, sand, aluminum or other metal turnings, etc., to facilitate distribution of the charge, or to employ perforated trays or special mechanical means for this purpose.

With respect to the multiple ash zones, each of which is in series with one of the reaction zones, the operating pressure may range from subatmospheric to about 200 p.s.i.g. It is preferred that each flash zone will function at a higher pressure than the next succeeding flash zone. Thus, the iirst hot ash zone will be at a reduced pressure below about 200 p.s.i.g.; the second flash zone at a reduced pressure below about p.s.i.g.; and, the third ash zone at a reduced pressure also below about 100 p.s.i.g., and also less than that of the second flash zone.

The catalytic composite disposed within the conversion zones which are in series with the ash zones, can be characterized as comprising a metallic` component having hydrogenation activity, and composited with a refractory inorganic oxide carrier material of either synthetic or natural origin. The precise composition and method of manufacturing the carrier material is not considered essential to the present process, although a siliceous carrier, such as 88.0% alumina and 12.0% silica, or 63.0% alumina and 37.0% silica, are generally preferred. Suitable metallic components having hydrogenation activity are those selected from the group consisting of the metal-s of Groups VI-B and VIII of the Periodic Table, as indicated in the Periodic Chart of the Elements, Fisher Scientific Company (1953). Thus, the catalytic composite may comprise one or more metallic components from the group of molybdenum, tungsten, chromium, iron, cobalt, nickel, platinum, palladium, iridium, osmium, rhodium, ruthenium, and mixtures thereof. The concentration of the catalytically active metallic component, or components, is primarily dependent upon the particular metal as well as the characteristics of the charge stock. For example, the metallic components of Group VI-B vare preferably present in an amount within the range of about 1.0% to about 20.0% by weight, the iron-group metals in an amount within the range of about 0.2% to about 10.0% by weight, whereas the platinum-group metals are preferably present in an amount within the range of about 0.1% to about 5.0% by weight, all of which are calculated as if the components existed within the finished catalyticcomposite as the elemental metal.

The refractory inorganic oxide carrier material may comprise alumina, silica, zirconia, magnesia, titania, boria, strontia, hafnia, and mixtures of two or more including silica-alumina, alumina-silica-boron phosphate, silica-zirconia, silica-magnesia, silica-titania, alumina-zirconia, alumina-magnesia, alumina-titania, magnesia-zirconia, titania-zirconia, magnesia-titania, silica-aluminazirconia, silica-alumina-magnesia, silica-alumina-titania, silica-magnesia-zirconia, silica-alumina-boria, etc. It is preferred to utilize a carrier material containing at least a portion of silica, and preferably a composite of alumina, silica and boron phosphate with alumina being in the greater proportion. The function of these rst reaction zones is essentially two-fold; they serve to concentrate a residuum fraction containing sulfur while simultaneously producing distillable hydrocarbons. The quantity of sulfur remaining in the distillable fraction will, of course, be dependent upon the characteristics of the fresh black oil charge stock. However, most of this remaining sulfur is concentrated in the middle-distillate and gas oil rangesfrom about 525 F. to about 1050 F.along with any residual nitrogenous compounds.

The last in the series of treating zones is utilized primarily to effect the virtually complete removal of nitrogenous and sulfurous compounds from the distillables, and to concentrate the residuum containing the exceptionally high-boiling contaminants. As hereinbefore set forth, the -successively higher operating severities in these reaction zones enhances the overall process by virtue of the fact that some hydrogenation and hydrocracking of the higher-boiling middle-distillate and gas oil components also takes place. The ranges of variable operating conditions in the final hydrocracking reaction zone include temperatures, as measured at the inlet to the catalyst bed, of from about 500 F. to about 850 F., pressures of 1,000 to about 4,000 p.s.i.g., hydrogen concentrations of from about 1,000 to about 50,000 s.c.f./bbl. and liquid hourly space velocities of from about 0.5 to about 5.0. Precise values for any one or more of these variables depends essentially upon two aspects which are necessarily considered in a given situation; (1) the nature of the distillable hydrocarbon mixture resulting from the first series of reaction zones, and (2) the ultimately desired product distributionie the relative proportions of gasoline, kerosene and middle-distillate boiling range hydrocarbons.

The effluent from the fourth, or hydrocracking zone, is introduced into a cold high-pressure separator functioning at substantially the pressure as the outlet of the conversion zone, and at a substantially lower temperature in the range of 60 F. to about 140 F. The gaseous phase, rich in hydrogen, is recycled at least in Apart to this hydrocracking zone, and, as hereinbefore stated, at least a part with both the first and second principally liquid phases. The liquid phase from the cold separator constitutes the product of the present invention, and may be subjected to standard fractionation and other separation techniques for the purpose of recovering specific select fractions.

It is preferred that the catalytic composite in the hydrocrackng zone comprise at least two refractory inorganic oxides, and preferably alumina and silica. When silica and alumina are employed in combination, the latter will be present within an amount of from about t0 about 90% by weight. Excellent results have been achieved through the utilization of the following silica-alumina composites: 88% by weight of silica and 12% by weight of alumina, 63% by weight of alumina and 37% by weight of silica, 88% by weight of alumina and 12% by weight of silica. The total quantity of the catalytically active metallic components will lie within the range of from about 0.1% to about 20.0% by weight of the total catalyst. The Group VI-B metal, such as chromium, molybdenum, or tungsten, when utilized in the hydrocracking catalyst, is usually present in quantities within the range of from about 0.5% to about 20.0% by weight of the catalyst. The Group VIII metals, which may be divided into two subgroups, are present in an amount of from about 0.1% to about 10.0% by weight of the total catalyst. When an iron subgroup such as iron, cobalt, or nickel, is employed, it is present in an amount of from about 0.2% to about 10.0% by weight, and preferably from about 1.0% to about 6.0%, whereas the platinum-group metal such as platinum, palladium, iridium, rhodium, etc., is present in an amount within the range of from about 0.1% to about 5.0% by weight of the total catalyst. When the metallic component of the hydrocracking catalyst consists of both a Group VI-B and a Group VIII metal, it will contain metals of these groups in a ratio of from about 0.05 :1 to about 5.0:1 of the Group VIII metallic components to the Group VI-A metallic components. Suitable catalytic composites, for utilization in the hydrocracking zone, comprise the following, but not by way of limitation: 6.0% by weight of nickel and 0.2% by weight of molybdenum; 1.8% by weight of nickel and 16.0% by weight of molybdenum; 6.0% by weight of nickel; 0.4% by weight of palladium; 6.0% by weight of nickel and 0.2% by weight of platinum; 6.0% by weight of nickel and 0.2% by weight of iron; 0.4% by weight of platinum; 6.0% by weight of nickel; 12.0% by weight of molybdenum; and, 6.0% by weight of nickel and 0.2% by weight of palladium.

In many instances, particularly when the catalytically active metallic components comprise metals selected from the platinum-group of Group VIII of the Periodic Table it will be desirable to include a halogen component to impart an additional acid-acting function to the hydrocracking catalyst. Such halogen is generally selected from the group of chlorine and/or iluorine, and will be present within the composite in an amount of from about 0.1% to about 8.0% by weight, calculated as the element, although referred to as combined halogen.

It is understood that the broad scope of the present invention is not to be unduly limited to the utilization of a particular catalyst having a particular concentration of components, a particular means for the manufacture of the same, or specic operating conditions other than those previously set forth. The utilization of any of the previously mentioned catalytic composites, at operating conditions varying within the limits hereinbefore set forth does not necessarily yield results equivalent to the utilization of other catalytic composites employed under other operating conditions.

The gaseous ammonia and hydrogen sulfide, resulting from the destructive removal of nitrogenous and sulfurous compounds, and light parafnic hydrocarbons, are removed from the total effluent in any suitable manner. For example, the efuent may be admixed with water, and thereafter subjected to separation such that the ammonia is absorbed in the water-phase. Hydrogen sulde and light parafnic hydrocarbons may be removed by introducing the effluent into a low-temperature flash chamber, the normally liquid hydrocarbons from which are passed into a fractionating column for the purpose of removing those hydrocarbons boiling within the gasoline boiling range. Other conditions and preferred operating techniques will be presented in conjunction with the following examples, one of which refers to the accompanying drawing in describing the applicability of our invention to a commercially-scaled unit.

EXAMPLES Recognition of the concept encompassed by the present invention was facilitated during the processing of the crude tower bottoms obtained from a sour Wyoming crude. The analysis of this charge stock is presented in Table I:

Table I.-Analysis of crude tower bottoms Gravity, API at 60 F. 11.4 Distillation, vol. percent at- 1000 F. 50.0 Heptane-insolubles, wt. percent 8.1 Sulfur, wt. percent 4.2 Vanadium, p.p.m. Nickel, p.p.m. 47

The tower bottoms was first processed over a catalytic composite of 2.0% by weight of nickel and 16.0% by weight of molybdenum composited with a carrier material of 68.0% alumina, 10.0% silica and 22.0% boron phosphate, by weight. The catalyst, about 64 grams, was disposed as a fixed-bed in a stainless steel tube having a nominal inside diameter of 7/a-inch; approximately one inch of quartz chips was placed above the catalyst to serve as a preheat zone. The operating conditions included a liquid hourly space velocity of 1.0, a pressure of 3,000 p.s.i.g., a hydrogen concentration of about 25,000 s.c.f./bbl. and a temperature gradient of 380 C. (at the top of the quartz chips) to 450 C. (at the bottom of the quartz chips). Following 153 hours of operation, it became necessary to pump methyl naphthalene through the system for the purpose of removing a deposit which had accumulated from the conversion product in the high pressure receiver. Up to this time, on the basis of product analyses, a deactivation trend with respect to sulfur removal and conversion of heptane-insolubles was indicated.

Following the naphthalene wash, the operation was continued at the same conditions with a slight improvement in hydrocracking and heptane-insoluble conversion being experienced. After 64 additional hours, it |was again necessary to wash the unit with methyl naphthalene for the purpose of removing the deposit from the high pressure separator. Another period of operation, at the same conditions, for a duration of about 40 hours, indicated decreasing hydrocracking activity accompanied by a higher residual heptane-insoluble content in the liquid product. At a catalyst life represented by a factor of 1.7 bbl/lb. (total quantity of charge stock per pound of catalyst disposed in the reaction zone), the ASTM distillation indicated 70.0% by volume of distillables at a temperature `of 1000 F. and having a gravity of 20.2 API at 60 F.

In a second operation utilizing the same charge stock and operating conditions with, however, the quartz chips being replaced by -30 mesh catalyst of the character hereinabove described. The operation was effected to a catalyst life, based only on the main body of catalyst, of approximately 4.6 bbl./lb. No plugging or pressure differential appeared with respect to the reaction zone, and no signicant deposit coated the high pressure receiver. At the termination of the operation, the product analysis indicated a gravity of 19.1 oAPI at 60 F. and 77.0% by volume of distillables was obtained at an ASTM distillation temperature of 1000 F. Although the gravity of the distillable product is somewhat less than that resulting from the previous operation, it must be noted that the second run extended more than three times as long and, more importantly, produced 7.0% more distillables, an increase of 10.0%.

In a third operation, the tower bottoms from the Wyoming Sour Crude was processed over 100 cc. of the nickel-molybdenum on alumina-silica-boron phosphate catalyst at a liquid hourly space velocity of 4.0. The pressure imposed upon the reaction zone was 3,000 p.s.i.g. and hydrogen was circulated by way of compressive means in an amount of 10,000 s.c.f./bbl. The temperature gradient within the reaction zone, the reactants flowing therethrough in downward flow, was controlled such that the catalyst bed inlet temperature was 370 C. (698 F.) while the outlet temperature was 400 C. (752 E)- At these conditions, the product analyses indicated the results shown in the following Table II under the designation of Run A.

The product from Run A was then processed in a second reaction zone at a liquid hourly space velocity of 2.0 and a temperature gradient of 420 C. (788 F.) at

the inlet to the catalyst bed and 470 C. (878 F.) at the outlet. Under these conditions, the product analyses indicated the results shown in Table II under the designation of Run B. A blend was prepared by admixing the total product from Run A, in equal parts by volume, with the 850 F.-plus portion of a single stage operation at the conditions of Run B. The synthetic charge stock was processed at a liquid hourly space velocity of 2.0 and a temperature gradient of 420 C. (788 F.) to 490 C. (914 F.). The results are presented in Table II under the designation of Run C.

The -foregoing results clearly indicate an improvement in hydrorefining/hydrocracking of black oil by (1) pretreating the black oil at a relatively low temperature insufficient to produce extensive cracking reactions, followed by processing at an elevated temperature; and (2) initially processing at a relatively low temperature, commingling the resulting product -with the heavier (850 F.-

plus in the example) portion, at least, of a single-stage, high temperature product, and processing the mixture at an elevated temperature. The present invention encompasses a method for processing black oils and which takes advantage of the concepts hereinabove described. In essence, the present invention involves processing a hydrocarbonaceous 'black oil in successively severe stages, in the presence of a portion of a previously processed productand Vwith intervening separation of at least a portion of the distillable material produced in each zone. Such distillable material is combined, following the final separation which concentrates the residuum fraction, with with the final step being the conversion thereof into lower-boiling hydrocarbon products.

In further describing our invention, reference will be made to the accompanying figure which is illustrative of one specific embodiment. In the accompanying drawing, the embodiment is presented by a simplified flow diagram in which details such as pumps, instrumentation, and controls, heat-exchange and heat-recovery circuits, valving, start-up lines, treating facilities and similar hardware have been omitted, or reduced in number, on the basis of not being necessary for an understanding of the inventive concept and the techniques involved therein. The use of such miscellaneous appurtenances, to modify the present process, are well within the purview of one skilled in the art of petroleum refining techniques.

For the purpose of demonstrating the illustrated embodiment, the drawing will be described in connection with the conversion of a black oil charge stock in a commercially scaled unit. It is understood that the char-ge stock, stream compositions, operating conditions, design and function of fractionators, separators and the like, are exemplary only, and may vary widely 4without departure from the spirit of my invention, the scope of which is defined in the appended claims. With reference now to the drawing, the charge stock is introduced into the process via line 1 in an amount of 10,000 bbl/day, being admixed therein with about 3,000 s.c.f./bbl. of a make-up hydrogen'stream in line 2 and a hydrocarbonaceous normally liquid recycle in line 3, the source of which is hereafter described. The black oil charge stock is a mixture of 70.4% by weight of a vacuum residuum, 16.0% by weightof a slop wax and about 13.6% by lweight of a catalytcally-cracked slurry oil. Inspection of this charge stock indicated a gravity of 10.0 API at 60 F., a Conradson Carbon `factor of 15.0 wt. percent, a sulfur concentration of 2.2% by weight, 3,260 p.p.m. of nitrogen and 4.1% by weight of heptane-insoluble asphaltenes. This commercially-scaled unit is designed to maximize the production of both a C6-400 F. end point gasoline fraction, and a 400 F. to 525 F. end point kerosene, with emphasis being placed on the former. Thus, as will be noted, fractionator 34 is indicated as providing a breakdown of the ultimate product into a Ca-minute overhead (line 35), a C4/C5 motor fuel blending component (line 36), a hexane-400 F. gasoline fraction (line 37), a 400 F.-525 F. kerosene cut (line 38) and a 525-650 F. gas oil fraction (line 39). For the present application, it will be assumed that the latter is recycled en toto through through line 3 containing valve 4 to combine with the fresh black oil charge stock to reactor 7.

In the unit being considered, the capacity is about 10,000 bbl./ day of black oil charge stock. In the interest of simplicity, the figure Will 'be described on the -basis of 100.0 mols/ hr. of such charge stock being introduced via line 1. The charge stock is admixed with sufficient hydrocarbon recycle in line 3 to provide a combined feed ratio to reactor 7 of 2.0; the liquid hourly space velocity based upon fresh feed only being of the order of 0.5. The mixture continues through line 3, is admixed with the make-up hydrogen from line 2, and passes therethrough into heater S wherein the temperature is raised to a level such that the temperature at the inlet of the catalyst bed in reactor 7 is 525 F.; the heated mixture passes through line 6 into reactor 7. The catalytic composite disposed in reactor 7 consists of 2.0% lby weight of nickel and 16.0% by weight of molybdenum combined with a carrier of alumina, silica and boron phosphate as hereinbefore described.

The first reaction zone eflluent, at a temperature of about 545 F., passes through line 8 into flash chamber 9 at a pressure of about 195 p.s.i.g., reduced from about 2,500 p.s.i.g. as imposed on reactor 7. In the following Table III, component analyses are presented for the stream in line 8 entering flash chamber 9 and the first vapor phase exiting therefrom via line 11. By way of a note of explanation, in Table III, and the following tables, the hydrogen is not shown nor is the quantity of liquid hydrocarbonaceous recycle through line 3.

Table IIL-Component analysis, first separation zone Lino Number 8 11 Component, wt. pcrcent'* *Propane and lighter normally gaseous components arc given in weight percent of charge as is customary.

It will be noted that very little conversion to lower boiling material is effected, although at least a portion o-f the sulfurous and nitrogen-ous compounds have been converted to yield ammonia and hydrogen sulfide. The normally liquid phase from flash chamber 9 passes through line 10, is admixed with a recycle hydrogen stream in line 40, and into heater 12 wherein the temperature is raised to a level of about 625 F.

The heated mixture enters reactor 14 via line 13 at a pressure of about 2650 p.s.i.g. The catalytic composite in reactor 14 is the same as that previously described with respect to reactor 7, and the amount is such that a liquid hourly space velocity of about 1.0 results in a temperature increase of about 75 F. to 100 F.-in the instant situation, the effluent leaves the reaction zone at about 700 F. The effluent passes through line 15 into flash chamber 16 at -a reduced pressure of about 75 p.s.i.g., and a second vapor phase leaves via line 18 to be combined with the first vapor phase in line 11. In the following Table IV, analyses are presented for the charge to reactor 14 in line 13 (the same composition yas the liquid phase in line 10), the eflluent in line 15 and the vapor in line 18.

Table IV.-Component analyses, second separation zone Line Number 13 15 18 Component, wt. percent:

Ammonia 0. 0. 10 Hydrogen Sultdem. 0. 75 0. 75 Methane.-. 0.50 0. 50 Ethane.. 0. 18 0.18 Propane 0. 31 0. 3l Component, mols/hr utane/Pentane 0. 9 0. 9 Hexane-400/ F 0. 3 3. 3 3.1 400 F.525 F. 0.5 3. 5 3. 2 525 F.650 F 1. 9 G. 9 6.0 650 F.1,000 F 96. 2 88. 3 1. 2

The liquid phase from flash chamber 16 is removed via line 17, combined wtih hydrogen recycle in line 41, and introduced into heater 19. The heated mixture, at a ternperature of 785 F. and a pressure of about 2700 p.s.i.g., passes through line 20 into reactor 21 containing the alumina-silica-boron phosphate catalyst containing 2.0% nickel and 16.0% molybdenum. The eflluent passes through line 22 into flash chamber` 23 at a reduced pressure of about 35 p.s.i.g., and the asphaltic bottoms Table V.-Component analyses, third separation zone Line Number 20 22 Component, wt. percent:

Ammonia,

Propane Component, mols/hr.:

Butane/Pentane Henne-400 F. 400 F.525 F.

99955-1 99.@ Odi-earn The mixture of the three vapor phases in line 11, following removal of the hydrogen sulfide and ammonia, is admixed with recycle hydrogen from line 42, the mixture continuing throfugh line 11 into heater 26. The heated mixture continues through line 27 into reactor 28 at a temperature of about 750 F. and a pressure of about 2950 p.s.i.g. Reactor 28 has disposed therein a catalytic composite of 5.0% by weight of nickel and `0.4% by lweight of palladium combined /with a carrier material of 63.0% alumina and 37.0% silica. The liquid hourly space velocity of 1.5 results in a catalyst bed outlet temperature of about 850 F. The reaction zone eflluent passes through line 29 into high pressure receiver 30, the latter at a temperature of about F. A hydrogen-rich -gaseous phase is refrnoved through line 31 via compressor 32. At least a portion continues and discharges through line 40 to combine with the material entering heater 12 and reactor 14; other portions are diverted through lines 41 and 42 to be combined with the charge to reactors 21 and 28 respectively.

Table VI.-Overall yields Wt. Percent Vol. Percent Componenti Ammoma weer@ Propane Butane/Pentan Hexane-400 F.

Residuum The liquid phase from high pressure separator 30 passes via line 33 into fractionator 34. Although illustrated as a single vessel for simplicity, it is understood that the separation of the final product may be effected by a combination of multiple distillations and/or fractionations. In any event, the light gaseous components are recovered as an overhead product from line 35, and a butane/ pentane fraction is removed Via line 36. The gasoline distillate is recovered as a hexane-400 F. fraction from line 37, and the kerosene cut, 400 F.525 F., is removed via line 38. The 525 F. to 650 F. material is removed as a bottoms fraction through line 39 and, in most situations not involving the desired production of a light gas oil, is diverted through line 3 containing valve 4, to combine with fresh hydrocarbon feed being introduced via linel. The amount of material diversion through line 3 is such that the combined, normally liquid feed ratio to reactor 6 is within the range of about 1.5 to 3.0. Where an excess of light gas oil is produced, it ymay be withdrawn as a valuable product through line 39, or stored for future use as liquid recycle should the need arise.

The foregoing specication and examples indicate the method by which our invention is put to use, and the results obtained thereby in the conversion of black oil charge stocks.

We claim as our invention:

1. A process for the conversion of a hydrocarbonaceous blac-k oil, containing hydrocarbons boiling above about 1050 F., into lower-boiling distillable hydrocarbons, Which process comprises the steps of (a) admixing said black oil with hydrogen and passing said mixture, at a temperature of from about 500 F. to about 600 F., into a rst catalytic conversion zone maintained at a pressure above about 1,000 p.s.1.g.;

(b) separating the resulting rst conversion zone eiuent in a rst separation zone at a pressure of from subatmospheric to about 200 p.s.i.g., to provide a rst vapor phase and a ilirst liquid phase;

(c) introducing said irst liquid phase into a second catalytic conversion zone, at a temperature above about 600 F. up to about 775 F., and a pressure above about 1,000 p.s.i.1g.; Y

(d) separating the resulting second conversion zone effluent in a second separation zone at a pressure of from subatmospheric to about 100 p.s.i.g., to provide a second vapor phase Iand a second liquid phase;

(e) introducing said second liquid phase into a third catalytic conversion zone, at a temperature above about 775 F. up to about 850 F., and a pressure above -about 1,000 p.s.i.g.;

(f) separating the resulting third conversion zone efliuent in a third separation zone at a pressure of from subatmospiheric to about 100 p.s.i.|g. to provide a third vapor phase and a third liquid phase containing hydrocarbonaceous material boiling above about 1050 F.;

(g) combining said first, second and third 'vapor phases with each other and with hydrogen, and introducing the resulting mixture into a fourth catalytic conwersion zone at a temperature of from about 500 F. to about 850 F. and a pressure of from about 1,000 to about 4,000 p.s.i.g.; and,

(h) separating the resulting fourth conversion zone eluent in a fourth separation zone at a temperature of from about F. to about 140 F. and a pressure substantially the same as imposed upon said fourth conversion zone, to provide a fourth liquid phase comprising distillable hydrocarbons and a fourth vapor phase rich in hydrogen.

References Cited UNITED STATES PATENTS 8/1933 Pier 208--59 7/ 1966 Inwood et al 208-59 

